WO2023083869A1

WO2023083869A1 - Process for eliminating troublesome secondary products in the direct oxidative esterification of methacrolein

Info

Publication number: WO2023083869A1
Application number: PCT/EP2022/081278
Authority: WO
Inventors: Steffen Krill
Original assignee: Röhm Gmbh
Priority date: 2021-11-12
Filing date: 2022-11-09
Publication date: 2023-05-19
Also published as: KR20240108464A; US20230150912A1; CN116120174A; US12017987B2

Abstract

The present invention relates to an optimized process for producing methyl methacrylate (MMA), wherein this MMA and polymers produced therefrom exhibit a very low yellowness index. According to the invention, the MMA is produced by direct oxidative esterification of methacrolein. In particular, the present invention relates to an optimized workup of the reactor discharge of the oxidative esterification of methacrolein, by means of which in particular coloring secondary products are removed or decomposed without having significant losses of MMA.

Description

Process for the elimination of interfering by-products in the direct oxidative esterification of methacrolein

field of invention

The present invention relates to an optimized process for the production of methyl methacrylate, this MMA and polymers produced from it being distinguished by a very low yellow value. According to the invention, the MMA is produced by direct oxidative esterification of methacrolein.

In particular, the present invention relates to an optimized processing of the reactor discharge from the oxidative esterification of methacrolein, by means of which particularly coloring by-products are removed or broken down without significant MMA losses.

State of the art

Methyl methacrylate is used in large quantities to produce polymers and copolymers with other polymerizable compounds. Furthermore, methyl methacrylate is an important building block for various special esters based on methacrylic acid (MAA), which can be produced by transesterification with the corresponding alcohol. This results in a great interest in manufacturing processes for this starting material that are as simple, economical and environmentally friendly as possible. It is of particular interest to provide methyl methacrylate (MMA) or other alkyl methacrylates which lead to slight yellowing during the polymerization.

MMA is now produced using a variety of processes that start from C ₂ - C ₃ - or C -building blocks. In a process that is considered to be particularly efficient, MMA is obtained by oxidizing isobutylene or tert-butanol with atmospheric oxygen in the gas phase at heterogeneous contact to form methacrolein and subsequent oxidative esterification reaction of methacrolein using methanol. This method, developed by ASAHI, is described, inter alia, in the publications US Pat. No. 5,969,178 and US Pat. No. 7,012,039. Disadvantage of this method is in particular a very high energy demand. In a further development of the process, the methacrolein is obtained from propanal and formaldehyde in the first stage. Such a method is described in WO 2014/170223.

US Pat. No. 5,969,178 describes such a process for the oxidative conversion of isobutene or tert-butanol to form methacrolein and the subsequent oxidative esterification to form MMA. In this second stage, a liquid mixture of methacrolein and methanol with a reduced water content is reacted with molecular oxygen and a palladium catalyst, which is usually present on a support as a palladium-lead catalyst. In a first distillation stage, a mixture of methacrolein and methanol is separated from the crude product of the oxidative esterification as overhead condensate of a rectification column, while low-boiling components, which contain acetone, for example, are removed overhead as vapors. The MMA-containing and low-methacrolein bottoms are then fed into a second distillation stage, with a saturated hydrocarbon, which is selected, for example, from hexane, heptane, octane and their isomers, being added to the top reflux of this second distillation stage. This saturated hydrocarbon forms an azeotrope with the methanol in the second distillation stage - for example the methanol/n-hexane azeotrope with a boiling point of 49.9 °C - whereby the methanol is separated from the MMA-containing phase that remains in the bottom. Furthermore, water also forms an azeotrope with the saturated hydrocarbon, such as the water/n-hexane azeotrope with a boiling point at 61.6° C., and is also separated from the MMA. A crude MMA remains in the bottom of the distillation column, which, in addition to the MMA, also contains other high-boiling products, such as, for example, methacrylic acid. At the top of the second distillation stage, the two azeotropes of methanol and water are condensed with the saturated hydrocarbon and fed into a phase separator. If required, fresh, saturated hydrocarbon can be added to this condensate stream in the form of a "make-up". The upper phase in the phase separator mainly contains the saturated hydrocarbon and is fed back to the second distillation stage as runback. The lower phase in the phase separator is passed into a third distillation stage, where a mixture of methanol and water is obtained as the bottom product. This mixture can then be fed to a dewatering between isobutene or tert-butanol oxidation to methacrolein and the actual direct oxidative esterification, as a result of which the amount of water in the direct oxidative esterification is reduced. The top product from the third distillation stage consists of the two azeotropes of methanol or water with the saturated hydrocarbon used and is returned to the phase separator in condensed form.

In this way, the educts of the direct oxidative esterification are recycled directly or indirectly and the raw MMA can then be purified to a sales quality.

Another alternative embodiment for recycling the starting materials is disclosed in US Pat. No. 5,969,178. The reaction product from the direct, oxidative esterification is passed into a first distillation column, with the feed to the column being located approximately in the middle of the column. In this distillation column, a mixture of, inter alia, methanol, methacrolein, MMA and water is obtained as top condensate. As before, low-boiling components which contain, for example, acetone can be removed overhead as vapors. The top condensate is a consequence of the two azeotropes of methacrolein and methanol with a boiling point of 58.0 °C and MMA and methanol with a boiling point of 64.5 °C. Analogously to the previous embodiment, this top distillate can then be fed into the dewatering between isobutene or tert-butanol oxidation to methacrolein and the actual direct oxidative esterification, but in this case the MMA must be recovered separately from the bottom product of the dewatering, otherwise the yield methacrylic product decreases.

In addition to MMA, the bottom product of the first distillation column also contains water and high boiler products such as methacrylic acid. After cooling, this bottom product is phase-separated and the organic, MMA-containing phase is subjected to conventional work-up to remove high and low-boiling products in order to obtain commercial MMA quality. The aqueous phase is discarded, resulting in a loss of yield in the form of MMA or methacrylic acid dissolved in water.

US Pat. No. 7,012,039 discloses a somewhat different work-up of the reactor effluent from the oxidative esterification. Here, in a first distillation stage, methacrolein is distilled off overhead via sieve trays and the aqueous, MMA-containing mixture is passed from the bottom into a phase separator. In this, the mixture is adjusted to a pH of about 2 by adding sulfuric acid. The separation of the sulfuric acid water from the organic or oil phase is then carried out using centrifugation. In a further distillation, this oil phase is separated into high-boiling components and a phase containing MMA, which is drawn off overhead. The MMA-containing phase is then separated from low-boiling components in a third distillation. This is followed by a fourth distillation for final purification.

The problem with this process is the sulfuric acid, which has to be added in large quantities and can have a corrosive effect on parts of the system. Accordingly, these parts, such as in particular the phase separator, which is described as a centrifuge, or the second distillation column, must be made of materials suitable for this purpose. Furthermore, US Pat. No. 7,012,039 is also silent about the handling of the methacrylic acid that is produced or the remaining methanol in the product. However, it can be assumed that the former is also separated in the distillation stages, while the methanol can only be partially recovered and recycled with the methacrolein, while the remainder is probably lost in the third distillation stage. The process per se is problematic with regard to the production of on-spec and colorless MMA, in particular through the oxidation of isobutene to methacrolein as the first process step, since significant proportions of diacetyl are produced there in the gas phase; Diacetyl is a well-known yellowing agent for MMA and PMMA made from it, the trace concentration of which must be meticulously controlled. Furthermore, there is no indication in this prior art of the amounts in which acetals of methacrolein, in particular of methacrolein dimethyl acetal, are formed and how these are treated and removed physically or chemically in the process. It should be noted that the addition of sulfuric acid in this patent serves in particular to separate lead cations in the form of poorly soluble lead sulphate in order to obtain an effluent which satisfies the requirements of municipal sewage water. The lead cations are leached from the second stage catalyst and lead acetate must also be continuously fed to the second stage to maintain catalyst activity.

WO 2014/170223 describes a similar method to US 7,012,039. The only difference is that in the actual reaction, the pH value is adjusted by adding a methanolic sodium hydroxide solution in a circuit. This serves, among other things, to protect the catalytic converter. Furthermore, the separation of the aqueous phase in the phase separation is easier due to the salt content. It leads but also to the fact that the resulting methacrylic acid is present as a sodium salt and is later separated and discarded with the aqueous phase. In the case of the variant of adding sulfuric acid in the phase separation, the free acid is indeed recovered. However, sodium (hydrogen) sulfate accumulates, which can lead to other problems during disposal.

Finally, EP 33 501 53 B1 (WO 2017/046110) teaches an optimized processing of the crude MMA obtained from an oxidative esterification, wherein the crude MMA is first separated from a heavy phase and then an alcohol-containing light phase is distilled off from this heavy phase , which in turn can be recycled. What is also special about this process is that the methacrolein was obtained on the basis of propanal and formaldehyde, with the former being obtained on the basis of C2 building blocks, e.g. from ethylene and synthesis gas. The patent teaches the processing of MMA-containing product mixture from a direct, oxidative esterification of methacrolein and methanol to MMA by unreacted educts, methacrolein and methanol, being drawn off overhead by distillation in a first distillation column and recycled into the reactor of the direct, oxidative esterification. The mixture of MMA, water and methanol obtained in the bottom also contains free methacrylic acid and the corresponding alkali metal salt and small amounts of methacrolein dimethyl acetal. This mixture is acidified and additional water is added.

A phase separation then takes place in a phase separator, followed by an extraction of the organic, upper phase. In this sequence of work-up steps, it is mentioned that the corresponding acetal is also split. However, no information is given as to how complete the cleavage of the acetal is and it is left open how the cleavage products methacrolein and methanol are worked up directly or indirectly. In addition, this teaching leaves open how much of the acetal remains in the final pure MMA and what impact this has on product quality for applications.

Overall, however, all of these processes, regardless of the raw material basis for the methacrolein used, lead to MMA or, in general, alkyl methacrylates, which lead to yellowing, in particular of the downstream products, such as PMMA molding compositions. There is therefore a need for improvement such as to remove the source of this yellowing as efficiently as possible from the corresponding alkyl methacrylate, in particular MMA, before the polymerization. In EP 34 504 22 A1, two different substances were identified as being relevant for this yellowing of the secondary products of an MMA produced by means of direct oxidative esterification. This is primarily a methacrolein dimethyl acetal (also referred to below as DMIB). In addition, methyl isobutyrate formed as a by-product in the process also appears to play a role in discoloration. The concept was now proposed to further convert the reactor discharge from the oxidative esterification in another reactor with reduced alcohol content, increased water content and lower pH. However, this method requires an additional reaction device. In addition, this reactor must be designed to be relatively large in order to achieve a residence time sufficient for the relevant reduction in by-products. The patent application teaches the treatment of a product mixture from the direct, oxidative esterification, with initially unreacted methacrolein and methanol being removed by distillation and the remaining bottom fraction containing MMA, water, methanol, free methacrylic acid and the corresponding alkali metal salt and small amounts of methacrolein-di -methyl acetal, is contacted with additional water and acid in an additional reactor, whereby substantial hydrolysis of the acetal is achieved. The acetal hydrolysis and the hydrolysis of methyl isobutyrate are described in various embodiments, such as a continuous stirred tank, a plug-flow reactor, a phase separator, or a column bottom. While the hydrolysis of the acetal and the methyl isobutyrate is quite efficient and an acetal content of less than 20 ppm is achieved, the application does not explicitly mention that the hydrolysis of MMA, and thus the valuable product of the process, also takes place under the conditions described. The yield of the target product thus falls and additional methacrylic acid and the addition products of water or methanol to the conjugated double bond are formed. Above all, the formation of hydroxyisobutyric acid is critical, since this component has a highly corrosive effect.

Task

In view of the prior art, it is therefore the object of the present invention to provide a technically improved process for an oxidative esterification of methacrolein for the production of MMA which, when processed further, leads to a particularly low yellowing of the subsequent products. In particular, the present invention was based on the object of efficiently removing the by-product DMIB, which is produced during the oxidative esterification of methacrolein and in particular leads to yellowing of the secondary products, from the MMA.

In addition, the task of keeping the content of the other by-product methyl isobutyrate in the final MMA as low as possible should also be solved.

The task was also to avoid by-products that lead to yellowing in such a way that the loss of MMA, in particular e.g. in the form of MMA hydrolysis to methacrylic acid, is as low as possible.

Furthermore, there was also the task of making available a method that can be operated with the least possible disposal effort, in particular due to a reduced accumulation of organic components and acids in the waste stream.

Solution

These problems were solved by developing a new method for producing an alkyl methacrylate, in particular MMA. This method has the following method steps: a. Production of methacrolein in a reactor I, b. oxidative esterification of methacrolein in the presence of an alcohol, oxygen and a heterogeneous catalyst containing noble metals in at least one reactor II c. Distillation of the reactor discharge from reactor II in a distillation column I to remove methacrolein and parts of the methanol, d. optionally a pretreatment of the bottom discharge from process step c comprising steps i. adding an acid to change the pH, ii. adding water to obtain a two-phase mixture and iii. reactively treating at least part of the mixture of ii, and e. an extraction of the bottom from distillation column I or at least one organic phase of the bottom discharge from distillation column I in an extraction I.

The invention is characterized in particular in that the amount of methacrolein in the organic discharge from extraction I is greater than the amount of methacrolein in the bottom discharge from distillation column I. According to the invention, this is achieved in particular by the hydrolytic conversion of methacrolein-based acetals and hemiacetals. This in turn is preferably realized by carrying out the reaction by means of a reactive distillation and/or reactive extraction.

The invention is additionally characterized in that the amount of methacrylic acid in the organic output from extraction I differs by less than 10%, preferably less than 5%, from the amount of methacrylic acid in the bottom output of distillation column I, with this being due to the methacrylic acid and its salts relates to the bottom discharge of the distillation column I. This is achieved by combining acetal cleavage in a reactor designed for this purpose with an extraction apparatus in which hydrolysis also takes place.

According to the prior art, the direct oxidative esterification of methacrolein to MMA proceeds with a selectivity to MMA of greater than 90%. In addition, depending on the conditions, such as the water content in the direct oxidative esterification, a selectivity to methacrylic acid of between 1% and 5% achieved, and a selectivity to methyl 3-methoxyisobutyrate between 0.5% and 3% found. Surprisingly, the process according to the invention permits a hydrolysis of the DMIB of more than 90%, but at the same time it allows the undesired hydrolysis of the MMA to form methacrylic acid to be reduced to a hydrolysis conversion of less than 3%.

The undesired hydrolysis can preferably be reduced to a conversion of less than 1%, very particularly preferably to a conversion of less than 0.5%.

The process according to the invention is preferably characterized in that distillation column I and extraction I are each designed to be acid-resistant, and that an organic and/or mineral acid is passed into the reactor stream from reactor II and/or into the bottom of distillation column I. This added acid preferably has a pKa value which is 1, preferably at least 2, lower than that of methacrylic acid.

In a first embodiment of the invention, the bottom discharge from distillation column I is passed in its entirety into extraction I, preferably into the bottom of extraction I. In this case, the acid is added, with liberation of methacrylic acid from the alkaline (earth) methacrylate, preferably either already in the distillation column I or in the bottom effluent of the distillation column I.

In a second, alternative embodiment of the present invention, the bottom discharge from distillation column I is first passed into a phase separator I, in which a separation into an organic, methacrolein-rich phase, which is in particular a phase consisting predominantly of MMA, and a polar, water-rich phase phase takes place. The phase separation is facilitated by the addition of water and the addition of an acid, with the bottom discharge from the distillation column undergoing a pH change. In this embodiment, the organic phase is then fed into the extraction I, while the polar phase is transferred to a distillation column IV. Alcohol fractions, in particular methanol fractions, can be separated from the aqueous phase in this distillation column IV. The alcohol-containing overhead fraction can then be recycled directly or indirectly into reactor II, while the aqueous bottom fraction can be disposed of. To improve the total consumption of raw materials, the bottom fraction from extraction I can also be fed into this distillation column IV.

In a variant I of the first embodiment, there is a reactor III between the distillation column I and the extraction I, in which the bottom discharge from the distillation column I is thermally treated with the addition of water before it is fed into the extraction I. This reactor III has the task of bringing about hydrolytic cleavage of methacrolein-containing acetals and hemiacetals.

In a variant II of the second embodiment, a reactor III is located between the distillation column I and the phase separator I, in which the bottom discharge from the distillation column I is thermally treated with the addition of water. This reactor III has the task of bringing about hydrolytic cleavage of methacrolein-containing acetals and hemiacetals. In all four embodiments or variants, the acid can be fed to the bottom stream from the distillation column I before it is introduced into the extraction I. In variants I or II, this takes place, for example, before reactor III or directly in reactor III. In the second embodiment, it is preferred to conduct the acid upstream of the phase separator I and/or directly into the extraction I. This effectively brings about the release of methacrylic acid from the respective (earth)alkaline methacrylate and at the same time the formation of the corresponding salt pair. When using sulfuric acid and sodium methacrylate, for example, sodium (hydrogen) sulfate is formed as a corresponding salt pair.

In the embodiments or variants described above, the acid is particularly preferably added before the phase separator I and optionally before the reactor III. The addition of the acid at these described positions has the advantage that the (earth) alkali metal salts from reactor II are preformed and are mostly present in the polar phase, as a result of which the subsequent phase separation and extraction I proceed more efficiently. Within these embodiments or variants, addition of the acid at more than one position is conceivable.

Furthermore, in the process according to the invention, the reactor discharge in distillation column I preferably has an average residence time of between 1 and 30 minutes, in particular between 1 and 20 minutes, at a temperature between 30 and 100.degree. In addition, the bottoms transferred from distillation column I, together with the acid transferred at the same time, remain in extraction I at a temperature of between 10 and 100° C., preferably at most 90° C. and particularly preferably between 20 and 60° C., for 1 to 20 minutes.

Furthermore, in the process according to the invention, reactor III preferably has an average residence time of between 10 seconds and 30 minutes, in particular between 1 and 5 minutes at a temperature of between 20 and 100° C., with the reaction mixture in reactor III preferably being turbulently mixed.

In particular, there are several alternatives with regard to the supply of the reactor stream mixed with an organic and/or mineral acid from reactor II. It is also possible, albeit less preferred, to feed the acid directly into the distillation column I, which is then mandatory in this embodiment designed to be acid-resistant. Alternatively, the mixture can also be mixed with another stream before it is fed into distillation column I. Furthermore, the stream can also be fed into extraction I or previously mixed with another stream before being fed into extraction I. Alternatively, the acid can also be fed into the feed of reactor III or directly into reactor III. It is also possible, although not preferred, to pass the acid and the reactor discharge separately into the distillation column I or into the extraction I.

It is important to note that the individual process steps do not have to be carried out continuously, directly one after the other. Further process steps, such as intermediate cleaning, in particular between the listed steps a and b. be performed. Process steps a to e are carried out, optionally supplemented by intermediate steps, in succession in the order given and in continuous operation. As a rule and preferably, steps b to d are carried out without such intermediate steps, while the methacrolein (MAL) from process step a. before use in process step b. first cleaned and drained.

The reaction in reactor II preferably takes place at a water content of between 0.1 and 10% by weight and a pH of between 5 and 8.

In principle, the methacrolein in process step a. in Reactor I based on C2 or C4 building blocks. Process step a is preferably the reaction of propanal with formaldehyde in the presence of at least one acid and optionally an amine, ie a process step which starts from C2 building blocks. In particular, the inventive method can be applied to the combinations of such a C2-based method for preparing the methacrolein and a subsequent oxidative esterification to form an alkyl methacrylate in method step b. This relates in particular to the descriptions for such a combination of method steps a and b as can be found, for example, in DE 3,213,681, US Pat.

Process step a in reactor I is preferably the reaction of propanal with formaldehyde in the presence of at least one acid and optionally one amine. For this purpose, the propanal is particularly preferably obtained from a C2 process. According to the invention, the C2 process describes processes which start from a C2 building block in the synthesis of an alkyl methacrylate. In the context of the present invention, particular preference is given to obtaining the propanal for process step a) on the basis of ethylene and synthesis gas.

Furthermore, the alcohol in process step b is preferably methanol and the alkyl methacrylate is MMA.

The addition of the acid splits methacrolein dimethyl acetal (dimethoxyisobutene, DMIB) into methacrolein and methanol in process steps d and e, or in variants I or II in reactor III and in process steps d and e with water. These cleavage products correspond to the starting materials used in process step b and, together with these remaining starting materials, are largely separated off overhead or in a side stream of the distillation column I and/or the extraction I and/or the optional distillation column IV and optionally directly or indirectly returned to Reactor II out. The phrase “to a large extent” usually includes at least 70% of the MAL and the methanol. If the recycling rate is calculated from the outlet of reactor II, this is preferably even more than 98% for MAL and methanol.

Alternatively, the top product from extraction I can also be separated by distillation into an MMA-containing phase and a low-boiling phase, with the starting materials from process step b being in the low-boiling phase. This low boiler phase in turn breaks down into an organic and an aqueous phase. Portions of the aqueous phase and/or the organic phase contain components that are difficult to separate from MMA, such as methyl isobutyrate or methyl propionate, and can therefore optionally be discharged from the process instead of being fully recycled into the process.

The acid can be added in process step d in such a way that it is passed directly into the bottom of the distillation column I. Alternatively or less preferably additionally, however, the acid can also be fed to the feed line from, for example, reactor II. It is also possible for the crude product from reactor II to be provided with the acid first in a mixing chamber before this mixture is fed into the distillation column I. When adding the acid to the bottom of a column, the combination of column design and choice of acid must be taken into account. Especially in the event that the column is filled to a relevant extent with distillation packing, no acid should be selected that forms a poorly soluble salt with potentially present ions, such as sodium ions. In order to prevent the column from clogging or clogging associated with this over time, methanesulfonic acid, for example, can be used instead of sulfuric acid if the acid is added to the bottom of the column.

The acid used is particularly preferably sulfuric acid. Sulfuric acid includes concentrated sulfuric acid, but also aqueous solutions of sulfuric acid. The use of an aqueous sulfuric acid with a sulfuric acid content of at least 10% by weight is preferred. Furthermore, the distillation in distillation column I is preferably carried out at a bottom temperature of between 50 and 100.degree. This internal temperature, measured in the liquid phase, depends in particular on the internal pressure and on the precise design of the distillation column used. Columns with intermediate trays, packing material or a larger internal volume can generally be operated at different temperatures than columns that do not have the respective configurations.

Optionally, additional water can be introduced into the bottom of the distillation column I to improve the results.

The acid is particularly preferably added to the outflow from the bottom of the distillation column I.

In an optional embodiment of the present invention, the acid, preferably sulfuric acid, can also be fed into the extraction in process step e, in which case the sulfuric acid can be concentrated, but aqueous solutions of sulfuric acid can also be used; the use of an aqueous sulfuric acid with a sulfuric acid content is preferred of at least 10% by weight. Depending on the design of the overall process, this additional acid in the extraction can result in a further reduction in the by-product content, in particular the DMIB. With a few modifications, the person skilled in the art can easily determine whether, in the chosen embodiment, an additional addition of acid in extraction I has a positive effect, or whether the acid carried over from process step d in extraction I is sufficient. Irrespective of this, method step e can be regarded as a reactive extraction. The cleavage of the methacrolein-containing acetals by means of water in an acidic environment is carried out using the process according to the invention in such a way that, after isolation, MMA can be achieved in a commercial quality with DMIB contents below 90 ppm.

The extraction I in process step d is preferably a mixer-settler. Alternatively, however, other devices such as, for example, a centrifugal extraction or a stirred cell extraction can also be used. The combination of a PlugFlow reactor as reactor III and a settler as phase separator I before extraction I has also proven to be very favorable. Optionally, additional water can also be fed in with one of these variants.

The medium used as extractant is preferably water, aqueous sulfuric acid or at least a proportion of the bottom stream from distillation column IV.

As a rule, process step e is followed by isolation and purification steps. These cleaning stages can be phase separators, distillation columns such as high and low boiler columns or crystallization chambers. It is preferably optionally at least one phase separator, at least one high boiler column and at least one low boiler column.

In process step e, in extraction I, the bottom discharge from distillation column I is separated into a phase containing alkyl methacrylate, in particular MMA, and a polar phase containing the acid used and salts thereof. On the other hand, formed methacrylic acid or other light organic acids can remain predominantly in the MMA-containing phase. The MMA-containing phase is then preferably passed into a distillation column II. This distillation column II is preferably a distillation column for removing high-boiling components from the MMA phase.

This distillation column II is particularly preferably designed to be acid-resistant. In an alternative embodiment, the distillation column II is a distillation column for separating off low-boiling components from the MMA phase. In this case, an acid, preferably sulfuric acid, in particular aqueous sulfuric acid, can be introduced into its bottom and/or feed line. In this embodiment, the bottom of this distillation column II is then passed into an acid-resistant designed distillation column III for the separation of high-boiling components from the MMA phase, so that in this particular embodiment the MMA-containing stream in four different process steps with the acid at a higher temperature, each time further reducing the level of by-products.

However, a further embodiment is very particularly preferred which is designed in the opposite way to the preceding one and in which the distillation column II is a high boiler column and the distillation column III is a low boiler column. In this embodiment, as a rule, no acid is added to the distillation column III.

According to the invention, the present invention is characterized in particular by the fact that in process step c in the bottom of the distillation column I the absolute water content is at least 2% by weight, preferably at least 3% by weight, particularly preferably at least 4% by weight higher than in reactor II In some embodiments of the present invention, in particular when the acid goes directly into the bottom of extraction I, the absolute water content difference can also be 20% by weight or even more than 30% by weight higher than in reactor II.

In addition, the alcohol concentration in process step c in the bottom of the distillation column I is lower than in reactor II in which process step b is carried out. This results from the distillation alone. However, it also seems to play a crucial role in relation to the decomposition of the DMIB.

Finally, this new method is characterized in that the pH value in the bottom discharge of the distillation column I after addition of the acid is generally between 0.5 and 7, preferably between 0.5 and 4, in particular between 0.5 and 3 is at least 0.5 lower than that set in reactor II. After the acid has been added, the pH values in extraction I are generally between 0.5 and 7. The acid can be added according to embodiments 1 and 2 and their variants I and II. The pH is preferably between 0.5 and 4, in particular between 0.5 and 3, and is preferably at least 0.5 lower than in reactor II.

Surprisingly, it was found that the alkyl methacrylates produced according to the invention, in particular MMA produced according to the invention, lead to particularly color-free end products. For example, in polymerization processes, the products are often subjected to degassing and the recyclates obtained through this degassing are returned to the polymerization in order to be used in a further polymerization step. This can lead to an accumulation of by-products, especially those with a lower vapor pressure during the process. This enrichment, in particular of dimethoxyisobutene (DMIB) and methyl isobutyrate, leads to an increasing yellowing of the polymer in later batches. Therefore, after a few recyclates, parts of the recyclate have to be discarded in order to deplete the recyclate cycle with regard to these by-products. This in turn leads to a reduction in the overall polymer yield. With the method according to the invention, it is now surprisingly possible to carry out significantly more batches with recycling and reuse of the recyclate and thus to significantly increase the overall polymer yield.

Surprisingly, it was also found that with the process according to the invention not only is the DMIB content reduced, but the proportion of undesired methyl isobutyrate in the end product is also reduced.

A surprisingly identified negative effect caused by the DMIB in an alkyl methacrylate resin and the secondary products produced therefrom is, in addition to the yellowing, reduced thermal stability of the polymer produced from the alkyl methacrylate. This is due to greater degradation of the polymer chain during thermal processing and is particularly evident when processing to form a shaped body and when processing a polymer syrup. A further advantage of the present invention is that an alkyl methacrylate produced according to the invention does not have these disadvantages in relation to the polymers produced from it. In addition to the process listed, novel alkyl methacrylates, which can be obtained, for example, as a product from process steps a to e according to the invention, are also part of the present invention. These novel alkyl methacrylates are distinguished by the fact that the alkyl methacrylate must contain DMIB as a component. In addition, the alkyl methacrylate usually also has methyl isobutyrate.

In particular, these alkyl methacrylates are those which can be prepared by means of a very advantageous process which starts from a C2 instead of a C3 or C4 base as the basic building block of the methacrylate synthesis. What is new about this alkyl methacrylate in particular is that although it has DMIB compared to the materials described in the prior art, this is present in an unknown content of less than 90 ppm, preferably less than 50 ppm, very particularly preferably less than 20 ppm. Contents of less than 90 ppm are particularly suitable for producing methacrylate resins without any visible yellowing. The alkyl methacrylate according to the invention also preferably has a methyl isobutyrate content of less than 250 ppm, particularly preferably less than 200 ppm and particularly preferably less than 170 ppm.

designation list

1 - Reactor II for the direct oxidative esterification of MAL and methanol to MMA

2 - optional feed of fresh MAL; Optionally, MAL can also be fed directly into 1

3 - possible supply of sulfuric acid

4 - Feed stream to distillation column I

5 - Distillation column I for recycling MAL and methanol

6 - recyclate stream or distillate stream from 5 with MAL and methanol to 1 (reactor II)

7 - possible supply of sulfuric acid and water

8 - sump drain from 5

9 - optional reactor III for splitting the acetal

10 - possible supply of sulfuric acid and water

11 - Flow from reactor III into phase separator I

12 - optional phase separators I

13 - Effluent from phase separator I - aqueous phase (salt rich)

14 - bottom stream of distillation column IV 15 - Distillation column IV, among other things for the recovery of further methanol

16 - distillate stream of the distillation column IV - can optionally be returned to reactor II or to 5 or to 6

17 - aqueous bottom effluent from extraction column I for feeding into distillation column IV

18 - effluent from phase separator I - organic phase (MMA rich)

19 - Extraction column I, among other things, for splitting DMIB (methacrolein dimethyl acetal)

20 - Addition of extraction water to the top of extraction column I.

Extraction water can come proportionately from 14

21 - Crude MMA, top effluent from extraction column I, is fed to distillation column II

examples

Example 1 - Cleavage of the Acetal

The direct oxidative esterification of methacrolein and methanol to form MMA was carried out in a stirred tank on an Au-cobalt oxide@Si02-AI203-MgO catalyst (see WO2017084969A1) at 80° C., 5 bar absolute, pH=7 and a molar ratio of methanol to methacrolein operated from 4 to 1 at the reactor inlet.

The reaction mixture obtained was fed to a distillation column (distillation column I) together with the fresh methacrolein; the pressure of the column at the top was 1000 mbar absolute, the reflux ratio was 1.00 and the bottom temperature was 73° C. on average. On average, the feed to the column was just under 2800 g/h and the bottom discharge was around 1180 g/h.

The relevant component concentrations of the reaction product (RKP) and the bottom of the distillation column I were:

All data in % by weight. An aqueous sulfuric acid solution was continuously added to the bottom discharge in a small, acid-resistant stirred tank, so that the pH was 2 and the ratio of bottoms to water feed was 0.77 (weight basis). The two-phase mixture was turbulently mixed in the stirred tank at 40° C. with an average residence time of 5 minutes. The 2-phase mixture was then phase separated at 40 °C and the two phases analyzed. The organic phase flow averaged 773 g/h and the aqueous phase flow averaged 1622 g/h.

The relevant constituent concentrations of the phases were:

All data in % by weight. | n.d. = not detectable. The detection limit is less than 5 ppm using GC-FID.

In addition, the MMA content in the organic phase of the phase separator was analyzed and showed an (absolute) decrease of 0.4%, while the total methacrylic acid content increased by (absolute) 0.3%. In total, this refers to the increase in methacrylic acid in the organic phase, but also in the aqueous phase of the phase separator, with the content being determined by means of HPLC in order to rule out that there is a lower result due to remaining methacrylic acid.

The organic phase was then run into the bottom of an extraction column, the extraction being carried out in a countercurrent process as a stirred cell extraction. The extraction parameters were analogous to the previous reactor 40°C and the pH was 2. The ratio of organics to extraction water was 3.9 (weight basis). The flow of crude MMA, i.e. the organics at the top of the extraction column, averaged 770 g/h.

The relevant component concentrations of the organics at the top of the extraction column (also called raw MMA) were:

The extraction water used was distilled water. It should be noted that the aqueous phase of the phase separator and the aqueous phase of the extraction can optionally be run together into a second distillation column in order to remove the remaining organics there by distillation. The sulfuric acid bottom product obtained from this distillation column can be used, at least in part, for acidification in the stirred tank and as extraction water, as a result of which the requirement for fresh water and fresh sulfuric acid in the process can be reduced.

Based on the analyzes and flows, the following sales can be calculated:

Conversion of the acetal cleavage in the stirred tank = 96.5%

Conversion of the acetal cleavage in the extraction column = 3.5%

Total conversion of acetal cleavage = >99.99%

The release of methacrolein can also be determined on the basis of the analyzes and flows:

It can thus be seen that the proportion and the amount of methacrolein increases due to the acetal cleavage and - at least partially - can be recycled in the direct oxidative esterification. The crude MMA obtained was purified using a high-boiling distillation column and a low-boiling distillation column and restabilized to a sales quality in a final pure distillation column. The MMA quality obtained is shown in Table 1.

Example 2 - pH variation for acetal cleavage

Example 2a - pH = 3

The experiment was carried out analogously to Example 1, but the pH for the acetal cleavage was 3. The following conversions were achieved:

Conversion of the acetal cleavage in the stirred tank = 73.2%

Conversion of acetal cleavage in the extraction column = 24.7%

Total conversion of acetal cleavage = 97.9%

Example 2b - pH = 4

The experiment was carried out analogously to Example 1, but the pH for the acetal cleavage was 4. The following conversions were achieved:

Conversion of the acetal cleavage in the stirred tank = 31.0%

Conversion of acetal cleavage in the extraction column = 59.4%

Total conversion of acetal cleavage = 90.4%

Example 2c - pH = 5

The experiment was carried out analogously to Example 1, but the pH for the acetal cleavage was 5. The following conversions were achieved:

Conversion of the acetal cleavage in the stirred tank = 9.7%

Conversion of the acetal cleavage in the extraction column = 32.5%

Total conversion of acetal cleavage = 42.2%

Examples 1 and 2 showed that, at low pH values, most of the acetal cleavage takes place in the stirred tank, and that acetal contents of less than 20 ppm can thus be achieved. As the pH increases, the proportion of conversion shifts in the direction of extraction, overall, however, conversions of more than 97% cannot be achieved, which would be necessary for a pure MMA with an acetal content of less than 20 ppm.

The further MMA processing was carried out analogously to example 1 and the results are summarized in Table 1:

Table 1 - Overview of pure MMA quality

It can be seen that with an increasing pH value and thus decreasing conversion during acetal cleavage, there is an accumulation of acetal and methyl isobutyrate in the pure MMA - as a result of which the yellow value in the polymer increases during later processing and thus the optical quality decreases.

Example 3 - Addition of the acid to the bottom of the distillation column I

The aqueous sulfuric acid was fed into the bottom of the distillation column I with the same mass flow as in Example 1, resulting in a pH of 2. The mixing resulted from the thermally induced circulation in the bottom of the column. The residence time in the bottom of the column was about 10 minutes. The stirred tank was dispensed with and the bottom of the column was run directly into the phase separation. The achieved conversion of acetal cleavage at the bottom of the column was greater than 99%, but due to the longer residence time and the higher temperature, side reactions or polymerisation formed, which resulted in a brown-colored mulm phase between the organic and the aqueous phase separation Phase formed, which disturbed the level control of the phase separation and the mass flow in the extraction fluctuated. This results in a fluctuating MMA content in the raw MMA and aqueous product of the extraction, which made the general plant operation more susceptible and the final pure MMA quality fluctuated. It is therefore in principle possible to carry out the acetal cleavage in the column bottom with sulfuric acid, but this means that the subsequent phase separation or extraction for long-term operation must be equipped with a technical option for removing the mulm phase in order to enable stable operation; this also results in additional costs.

Example 4 tubular reactor instead of the stirred tank for acetal cleavage

A tubular reactor with a length of 60 m and an internal diameter of 4 mm and 3 static mixers evenly distributed over the reactor was flown with the bottom discharge of the first distillation column (for the composition see Example 1) and an aqueous sulfuric acid solution in such a way that the residence time in the reactor was 83 seconds. The ratio of organics to aqueous sulfuric acid was 0.71 (weight basis) and the pH was 3. The tubular reactor was operated at 80° C. by preheating the feed streams and jacket heating.

Downstream of the tube reactor, the reaction mixture was cooled to 40° C. and the phases were separated analogously to Example 1. The conversion of the acetal cleavage was determined by GC analysis and was 96.1%, which is higher than with the stirred tank. This example shows that improved mixing and a higher reaction temperature can increase the space-time yield of the acetal cleavage while at the same time reducing the consumption of sulfuric acid (tubular reactor = 1.37 mol DMIB / kg H ₂ SO ₄ x hr vs. stirred tank = 1.2 mol DMIB/kg H ₂ SO ₄ ×hr).

Example 5 Acetal cleavage in a stirred tank at elevated temperature

The test was carried out analogously to Example 1, but the temperature was increased from 40 to 80.degree. The following sales were achieved:

Conversion of the acetal cleavage in the stirred tank = 98.9%

Conversion of the acetal cleavage in the extraction column=1.0% Total conversion of the acetal cleavage=>99.99% However, the MMA content in the organic phase of the phase separator decreased by 2.8%. At the same time, the methacrylic acid content in the aqueous and organic phase of the phase separator increased by a total of 2.3%. In total, this refers to the increase in methacrylic acid in the organic phase, but also in the aqueous phase of the phase separator, with the content being determined by means of HPLC in order to rule out that there is a lower result due to remaining methacrylic acid.

This suggests that as the temperature rises, the MMA is hydrolyzed into methacrylic acid.

Example 6 Acetal Cleavage in a Stirred Tank at Elevated Temperature and Elevated Residence Time

The experiment was carried out analogously to Example 1, but the temperature was increased from 40 to 80° C. and the residence time was doubled from just under 5 to just under 10 minutes. The following sales were achieved:

Conversion of the acetal cleavage in the stirred tank = 99.8%

Conversion of acetal cleavage in the extraction column = 0.1% Total conversion of acetal cleavage = >99.99%

However, the MMA content in the organic phase of the phase separator decreased by 5.4%. At the same time, the methacrylic acid content in the aqueous and organic phase of the phase separator increased by a total of 4.5%. In total, this refers to the increase in methacrylic acid in the organic phase, but also in the aqueous phase of the phase separator, with the content being determined by means of HPLC in order to rule out that there is a lower result due to remaining methacrylic acid. This suggests that as the temperature rises, the MMA is hydrolyzed into methacrylic acid.

Examples 5 and 6 show that at elevated temperatures and/or residence times, in addition to an increased conversion of the acetal cleavage, there is also a noticeable hydrolysis of MMA, as a result of which valuable material is lost.

Example 7 Acetal cleavage in a stirred tank with varied pH The experiment was carried out analogously to example 1, but the pH for the acetal cleavage was 1. The following conversions were achieved:

Conversion of the acetal cleavage in the stirred tank = 99.8%

However, the MMA content in the organic phase of the phase separator decreased by 2.5%. At the same time, the methacrylic acid content in the aqueous and organic phase of the phase separator increased by a total of 2.1%. In total, this refers to the increase in methacrylic acid in the organic phase, but also in the aqueous phase of the phase separator, with the content being determined by means of HPLC in order to rule out that there is a lower result due to remaining methacrylic acid. This suggests that as the temperature rises, the MMA is hydrolyzed into methacrylic acid.

Example 7 shows that a sufficiently high conversion in the acetal cleavage can also be achieved without the extraction, but at the cost of MMA loss and thus loss of the valuable substance or product of the process.

Nevertheless, Examples 5 to 7 are examples according to the invention, since the proportion of hydrolyzed MMA is still lower here than in the prior art and at the same time the acetals are efficiently cleaved.

Claims

26

Expectations

1 . Process for the production of alkyl methacrylates, comprising the process steps a. Production of methacrolein in a reactor I, b. oxidative esterification of methacrolein in the presence of an alcohol, oxygen and a heterogeneous catalyst containing noble metals in at least one reactor II c. Distillation of the reactor discharge from reactor II in a distillation column I to remove methacrolein and parts of the methanol, d. optionally a pretreatment of the bottom discharge from process step c comprising steps i. adding an acid to change the pH, ii. adding water to obtain a two-phase mixture and iii. reactively treating at least part of the mixture of ii, and e. an extraction of the bottom from distillation column I or an organic phase of the bottom discharge from distillation column I in an extraction I, characterized in that the amount of methacrolein in the organic discharge from extraction I is greater than in the bottom discharge from distillation column I.

2. The method according to claim 1 for the production of alkyl methacrylates, comprising a method step e., which is an extraction of the bottom of distillation column I in an extraction I, characterized in that extraction I and optional distillation column I are designed to be acid-resistant, that an organic and/or a mineral acid is mixed with the reactor stream from reactor II, is passed into the distillation column I and/or the extraction I, that the reactor discharge has an average residence time in the distillation column I of between 1 and 30 min at a temperature of between 30 and 100 °C, and that in extraction I in the presence of that transferred from distillation column I Bottom with the additionally converted acid at a temperature between 10 and 100 ° C for 1 to 20 min in the extraction. Process according to Claim 1 or 2, characterized in that the reaction in reactor II takes place at a water content of between 0.1 and 10% by weight and a pH of between 5 and 8. Process according to at least one of Claims 1 to 3, characterized in that process step a in reactor I is the reaction of propanal with formaldehyde in the presence of at least one acid and optionally an amine, and that the alcohol in process step b is methanol and the alkyl methacrylate is MMA. The method according to at least one of claims 1 to 4, characterized in that in process steps c methacrolein dimethyl acetal is split with water into methacrolein and methanol, which in the distillation column I together with the remaining starting materials of process step b. be removed to a large extent overhead or in a side stream of the distillation column I and returned to reactor II. Process according to at least one of Claims 1 to 5, characterized in that in process step c sulfuric acid is fed into the distillation column I as said acid and optionally additionally in process step e into the extraction. Process according to at least one of Claims 1 to 6, characterized in that additional water is introduced into the bottom of the distillation column I and that the distillation takes place at a bottom temperature of between 50 and 100°C. Process according to at least one of Claims 1 to 7, characterized in that extraction I is a mixer-settler and the extraction agent used is water, aqueous sulfuric acid or the bottom stream from distillation column IV.

9. The method according to at least one of claims 1 to 8, characterized in that the acid is added to the effluent from the bottom of the distillation column I.

10. The method according to at least one of claims 1 to 9, characterized in that on method step e. Isolation and purification stages follow, which are at least one optional phase separator, at least one heavy ends column and at least one lights column.

11. The method according to claim 10, characterized in that in method step d. in extraction I, the bottom discharge from distillation column I is separated into an MMA-containing phase and a polar phase containing the acid used and its salts, that the MMA-containing phase is then fed into a distillation column II, which is preferably a distillation column for Separation of high-boiling components from the MMA phase is conducted.

12. The method according to claim 11, characterized in that the distillation column II is designed to be acid-resistant, and that an acid, preferably sulfuric acid, is introduced into the bottom and/or the feed line of the distillation column II.

13. The method according to claim 11, characterized in that it is in the distillation column II is a distillation column for separating low-boiling components from the MMA phase, in the bottom and / or feed an acid, preferably sulfuric acid is introduced, and that the The bottom of this distillation column II is passed into an acid-resistant distillation column III for separating high-boiling components from the MMA phase.

14. The method according to at least one of claims 1 to 13, characterized in that the acid added has a pKa value which is 1 less than that of methacrylic acid. 29 MMA, preparable by means of process steps a to d of a process according to at least one of claims 1 to 11, characterized in that the MMA has methacrolein dimethyl acetal (DMIB) and methyl isobutyrate and a DMIB content of less than 90 ppm and a methyl isobutyrate content less 200ppm. MMA according to Claim 15, characterized in that the MMA has a DMIB content of less than 50 ppm and a methyl isobutyrate content of less than 250 ppm.